Chemical vapor deposition (CVD) reactors are used to produce polycrystalline silicon (polysilicon), the key raw material used in the manufacture of most semiconductor devices and silicon-based solar wafers and cells. The most widely used method for producing polysilicon is the Siemens reactor process generally according to the primary reactions below:HSiCl3+H2→3HCl+SiHCl+HSiCl3→SiCl4+H2 Net: 4HSiCl3→Si+2H2+3SiCl4 
In commercially significant polysilicon processes, STC, H2, dichlorosilane (DCS) and HCl are all significant byproducts of the CVD operations along with lesser amounts of monochlorosilane (MCS), silane (SiH4), and trace contaminants containing metals, donor and acceptor atoms, and carbon bearing species. In addition to TCS, significant amounts of H2 are fed to the CVD reactor and the single pass conversion of TCS is well under 50%. Variations of this process have been in commercial existence for about fifty years and are widely reported in the literature. In this process, high temperature polysilicon rods are placed in a reactor, and trichlorosilane (TCS) gas is passed over these rods. A portion of the silicon in the gas is deposited on the rods, and when the rods have grown large enough, they are removed from the reactor. The end product is in the form of polysilicon rods or chunks, which can be further processed into ingots, then sliced into wafers that are made into solar cells, for example.
In a related process, TCS is disproportionated to form silane (SiH4) and STC. The silane produced is used in many processes associated with semiconductors and other products, including making polysilicon in either a Siemens reactor or fluidized bed CVD process. The fluidized bed process makes silicon in irregular, but nominally spherical beads in diameters typically ranging up to about 2 mm diameter. The general chemistry of these reactions is as follows:4HSiCl3→SiH4+3SiCl4 SiH4→Si+2H2 In the exhaust of a TCS CVD reactor, the chlorosilanes (Six—Cl4−x) are separated from the the non-condensable gases (H2 and HCl) by relatively simple condensation steps and then followed by energy intensive and capital intensive processes consisting of compression, HCl absorption, HCl stripping, and H2 pressure-swing adsorption (PSA) steps to remove HCl from the H2. The by-products of the STC converter are similar to those produced in the CVD reactors and have traditionally been subjected to the same or very similar off-gas recovery (OGR) designs.
In all commercial processes using TCS as a feedstock, the initial step of production of TCS can be further classified as direct chlorination process or the hydrochlorination process, both processes being widely referenced in the literature, though nomenclature is not necessarily consistent. It is the TCS thus produced that is reacted with H2 in Siemens style chemical vapor deposition (CVD) reactors or disproportionated to make SiH4.
The process converting TCS into silane and the CVD-based Siemens process for manufacturing polysilicon both produce a large amount of the byproduct silicon tetrachloride (STC). For example, a maximum of about 20 kg of STC is made as a byproduct for every kg of polysilicon or silane produced. It is possible, however, to hydrogenate STC forming TCS by reacting STC with hydrogen in the gas phase in a reactor commercially referred to as an STC Converter at 800-1200° C. where 14-24% of the STC is typically hydrogenated in each pass through the reactor according to the following reaction:SiCl4+H2→HCl+HSiCl3 The product TCS can then be recycled to a series of silane disproportionation reactors and separation steps to make silane, or to a CVD reactor for direct production of more polysilicon.
Most commercial STC converter processes use molar feed ratios of H2:STC between 2.0:1 and 3:1 according the process of FIG. 1. In this process, the STC converter reaction chemistry is insensitive to pressure, but the industrially practiced range has been predominantly 4-7 BarG. Most commercially available STC converters are retrofitted Siemens CVD reactors having multiple graphite rods for heaters and a flat baseplate where feed-through electrical connections are made to the graphite rods. When conducted in a Siemens style STC converter, it is costly to increase the pressure due to a large flat baseplate forming the bottom of the reactor and the reaction step itself does not benefit from increased pressure. In the entire STC converter process to make TCS, the OGR system which separates reactor effluent products and reactants and recycles unreacted feeds must be considered, and this part of the process benefits significantly from operation at higher pressure.
The OGR process is a very energy intensive process mainly due to the difficulties of separating H2 and HCl. Since, hydrogen and HCl cannot be separated by a simple condensation process and would require excessive energy if separated via cryogenic distillation, alternate methods such as membrane separation processes have been evaluated. Even a membrane separation process is not deemed feasible for separation of H2 and HCl. No durable membrane technology exists at a competitive capital cost. Consequently, polysilicon producers typically use STC to absorb HCl from the H2/HCl mixture. TCS, which has higher vapor pressure than STC, has also been used to absorb HCl. Whether using TCS or STC to absorb HCl from H2, significant amounts of the chlorosilane exist in the H2 when it is subsequently fed to a PSA (pressure swing adsorption) or TSA (temperature swing absorption) operation to remove residual HCL and trace impurities. The presence of these chlorosilanes reduces the capacity of PSA or TSA adsorbent beds to adsorb the impurities of concern. Use of either TCS or STC for HCl absorption is an energy intensive process.
Another process to hydrogenate STC to TCS is what is commercially known as a hydrochlorination FBR process. In this process, MGSI, H2, and STC are fed into a FBR where the following net reaction occurs, typically at a temperature of 500 to 600° C. and typically 15-30 BarG.Si+2H2+3SiCl4→4HSiCl4 In this process, approximately 15-25% of the STC fed is reacted in a single pass and typical molar feed ratios are 1.5:1 to 2.5:1 H2:STC. The process benefits significantly from higher pressures and temperatures due to improved equilibrium conversion. As a result, the process is very expensive due to high pressure and temperature ratings required on relatively large process equipments. Vessels must be made of fairly thick walls with high nickel alloys. Both the metals used and the fabrication techniques are quite expensive.
If STC could not be hydrogenated and recycled, there would be a huge loss of the primary raw materials silicon and chlorine and a cost for disposal of the byproduct STC. Thus, efficient polysilicon plants are built as a substantially closed loop processes as illustrated in FIGS. 1 and 2. FIG. 1A shows a typical polysilicon plant utilizing hydrochlorination technology. In this drawing, a hydrochlorination plant 2 provides TCS to CVD operations 1 and the CVD operations return H2 and STC to the TCS plant. Impurities and byproducts are purged from the TCS plant and make-up raw materials MGSI, H2, and STC/HCl are fed to the plant. The hydrochlorination plant produces TCS from MGSI and also hydrogenates STC from the CVD operations.
FIG. 1B shows a typical polysilicon plant utilizing a direct chlorination plant 4 to make TCS from MGSI plus HCl and an STC hydrogenation plant 3 to hydrogenate STC plus H2 to TCS. CVD operations 1 are the same as in the hydrochlorination based plant.
FIG. 1C shows a silane plant 5 coupled to a TCS plant 6, and CVD operations 7. In FIG. 1C, TCS plant 6 can represent the non-CVD operations of FIGS. 1A and 1B. In this FIG., TCS is supplied from the TCS plant to a silane plant which produces Silane and returns STC according the reaction:4HSiCl3→SiH4+3SiCl4 It should be understood that in FIGS. 1A, 1B, and 1C, many operations occur within each block and additional minor feed streams may exist as feeds enter and byproduct and impurity streams may leave the processes.
FIG. 2 shows a typical hydrochlorination synthesis plant. MGSI 242 is conveyed into a low pressure hopper 201 through line 248. A series of valves is manipulated in lines 248 and 202 in order to transfer the abrasive solid through line 202 into high pressure hopper 203. Valves are then manipulated to transfer the abrasive solid through line 204 into the hot and high pressure FBR 205. MGSI reacts with hot feed gases coming in through line 217 and FBR product leaves through line 206, where it may go through interchanger 207 and line 208, or go directly to quench vessel 209.
The primary function of the quench 209 is to stop fine particulates of MGSI and metals salts such as FeCl3 and AlCl3 from moving further through the process train. (Fe and Al are impurities typically present in MGSI that from volatile salts in the FBR). This is accomplished by countercurrent flow of a dilute slurry (lines 230, 231 and pump 233) of the fine particulates in STC and TCS through a packed bed against the rising vapors. A portion of the dilute slurry is fed through line 232 to stripper 234 where it flows counter current against a H2 stream 228 fed into the bottom of the stripper and leaving the top of the stripper as line 229 saturated with volatile components. This concentrates the silicon particulates that may have been carried out of the FBR 205 and also the volatile metal salts into a residue stream. The residue stream 235 comprising STC, TCS and solids slurry is finally discharged into a low pressure slurry handling and recovery unit 236 where some portion of the STC and TCS present are recovered via stream 237 and join the condensate stream 247 forming stream 223. Solids and residual chlorosilanes plus impurities in stream 238 are treated for final disposal in unit 239. A typical practice is to hydrolyze and neutralize stream 238 with neutralizing media 240, leaving waste stream 241 for disposal as dictated by local practices and regulations.
Vapor leaves the top of quench vessel 209 through line 210 where some of the chlorosilanes in the vapor stream are condensed in condenser 249. Some of the condensate 219 is refluxed via line 220 and some is drained via line 246 and ultimately fed to crude distillation column 244 where STC and TCS are separated. Vapors 211 leaving condenser 249 continue to a series of condensers represented by 250. Typically, the terminal temperature in the series of condensers 250 is −25 to −50 C. The H2 and relatively small concentrations of chlorosilanes leave condenser 250 in stream 212 where they are compressed in compressor 243, flow through line 213 to heater 247 and then join a stream of vaporized STC 227 coming from STC vaporizer 226. Stream 225 represents STC feed from other parts of the plant (CVD and/or silane operations) and stream 252 is a blowdown from the vaporizer.
TCS 245 leaving crude column 244 goes to further purification and from there to CVD operations 1. STC 224 leaves the bottom of the crude column and is fed to STC vaporizer 226 where it is vaporized and the mixed with H2 coming from heater 247. This combined stream 214 is the fed through interchanger 207 (if an interchanger is used) through line 215 to trim heater 216 and finally into the bottom of the FBR as stream 217.
From the description above, it is apparent that the process involves many steps. The combination of high temperatures, high pressures, abrasive solids and corrosive environment make the process equipment quite specialized and expensive. Valves in MGSI lockhoppers and associated lines used to charge the FBR wear out. The FBR and quench vessel are typically made of expensive high nickel alloys that require specialized fabrication techniques which are also expensive. Handling abrasive slurries in bottoms streams at high pressures requires specialized pumps and careful design. Even then, these components have wear limited lifespans.
Since temperatures over 500° C. are required to achieve desirable kinetics and equilibrium in the FBR, heating feed streams uses significant thermal energy. Electric heaters typically used for this service are expensive to purchase and/or have limited life. Heat integration from the FBR exhaust stream before the quench is significantly complicated and limited due to fouling from deposition of volatile chloride salts of iron and aluminum byproducts created in the FBR from their respective metallic impurities in the MGSI.
For the reasons above, the hydrochlorination process is a relatively expensive process to build and to operate. The size of equipment is largely dictated by the H2 flow through the FBR. Difficulties associated with availability of materials like metal plate and fabrication difficulties with the alloys used limit the size of FBR's and to a lesser extent quench vessels. Even with FBR's built as large as possible based on fabrication capabilities and process understanding, world scale plants today have as many as 8 or more parallel hydrochlorination trains. This is far more than are required for reliability. Multiple units are necessary primarily because the capability to build larger trains does not exist and/or the incremental cost of larger trains is prohibitive due to fabrication difficulties. It is reasonable to expect technology will evolve and larger units of substantially the same design can be built in the future, but these size increases using parallel technology are expected to be somewhat incremental and without significant impact on the economics. A solution increasing capacity and/or capital cost per unit capacity is quite valuable to the industry.
Depending on efficiency of the process, energy costs can be either the greatest or second greatest cost in producing TCS in a polysilicon plant. Purchased MGSI is typically the other highest cost though economics vary from site to site. Reducing energy requirements are key to reduce operating costs.
FIG. 3 shows detail for a typical STC hydrogenation plant 3 of FIG. 1B. In this process, STC 327 from CVD reactors or a silane plant (stream 326) and recycle stream 325 is fed to STC vaporizer 329 and then through line 301 where it is mixed with H2 from stream 302 and fed through line 303 into STC converter 304. The molar ratio of feeds is typically 2.0:1 to 3:1 H2:STC. In the STC converter approximately 18% of the STC fed is reacted to form TCS and HCL. The product stream 305 flows through a condensation train 306 consisting of one or more condensers and then through line 307 to a compressor 308 where the pressure is elevated to enable recycle of the H2 and also to make subsequent condensation more efficient. After leaving the compressor, the gas flows through line 309 to condensation train 310 representing one or more condensers. Stream 311 then flows into a countercurrent HCl Absorber 312 where cold chlorosilanes 313 are flowing countercurrent to the rising gas coming from 311. HCl and more volatile chlorosilanes are absorbed and exit through the bottom of the absorber via stream 314. Stream 314 then flows to HCl stripper column 315 where the HCl is distilled away from the chlorosilanes via stream 319 and is directed to the direct chlorination plant 4 of FIG. 1B. Chlorosilanes free of HCl leave the bottom of the HCl stripper in stream 316. Part of the flow follows line 321 to crude column 323 where STC and TCS are separated. TCS 324 is fed to a purification train and then to CVD operations 1 of FIG. 1. The remainder of the flow from the HCl stripper bottoms stream is fed through line 322 to absorber feed cooler 318 and then through line 313 into the HCl absorber. STC from the crude column is fed through line 325 where it is combined with STC 326 from CVD operations 1, and fed to the STC Vaporizer. A small blowdown 328 from the STC vaporizer 329 purges impurities.
It should be understood that variations of this flow path utilizing various heat integration schemes are possible, and that the essential steps of an STC converter OGR used in industry are nearly universal. Those steps being condensation, compression, HCl absorbtion and HCl stripping. Many plants include a PSA (pressure swing adsorption) and/or TSA (temperature swing adsorption) step where the H2 of stream 302 flows through a solid media to further adsorb HCl and other components from the H2. This TSA or PSA process consumes additional energy in that it must be heated and cooled (for TSA) or components are periodically purged into a low pressure stream (the PSA) process. The energy requirements in the TSA process are heating and cooling of the adsorptive media along with providing cooling and heating associated with the heat of adsorption and desorption. In the PSA process, either large amounts of purge gases are lost, or secondary recovery systems must be built to recover valuable gases from the purge.
The OGR in an STC hydrogenation process producing HCl byproduct is a significant capital investment and energy intensive part of the process. Refrigeration equipment is expensive and consumes large amounts of electricity to cool stream 313 feeding the HCl Absorber and run a condenser in HCl stripper 319. Thermal energy required to strip HCl from the chlorosilanes in the HCl Stripper is also quite large. Since most STC converters operate at 4-7 Bar pressure and the HCl Absorber typically runs at 12-15 bar, significant energy is consumed in the compressor. Due to fairly low suction pressure, the compressors required are relatively large and capital intensive as well. When efficient high capacity converters as described in U.S. Patent Publication No. 2012/0328503 A1 are used, the cost of the OGR significantly exceeds the cost of the STC converters and most of the cost in the OGR is in the portions dedicated to removing HCl from H2 and subsequently recovering the HCl in a relatively pure form. An OGR having lower capital and operating costs represents a high potential to reduce capital expenditure and energy consumption.
All patents, patent applications, provisional patent applications and publications referred to or cited herein, are incorporated by reference in their entirety to the extent they are not inconsistent with the teachings of the specification.